Hydrocarbon conversion



Mmh 23, 1941s.`

R. P. RUsAsELl. ETAI. 'HYDROCARBON CONVERSIONY Y5 sheets-sheet 1 Filed Aug. 2l, 1942 R. P. RUSSELL I'AL .HYDROCARBON CONVERSION March 23, 1948.

Filed Aug. 21, 11942 5 Sheets-Sheet 2 Mamzhy 23, 194s.

72ACT/OIVA TINC R. P. RUSSELL ET AL HYDROCARBON CONVERSION Filed Aug. 2l, 1942 VESSEL c D callan/Maron 5 Sheets-Sheet 3 HYopQcAReo/vs, onus THAN H Ya/rocmvaoA/s or Y New: mmv 4c-Arom March 23, 1948.

R. P. RUSSELL ETAL HYDROCARBON CONVERS ION 'WWW www@ March 23, 1948. R. P. RUSSELL' Erm. 25438456 HYDROCARBON CONVERSION Filed Aug.4 2l, 1942 5 Sheets-Sheet 5 M. qu iwan Patented Mar. 23,v 1948 HYDROCARBON CONVERSION Robert P. Russell, Short Hills, Eger V. Murplhree, Summit, and Charles E. Hemmingen', Westfield, N. J., assignors to Standard Oil Development Company, a corporation oi' Delaware Application Aug-ust 21, 1942, serial No. '455,574

e claims. (ci. 26o-683.4)

1 This invention relates to the conversion of hydrocarbon oils and involves an improved process for the production of high quality aviation gasoline and of aviation basenaphtha and alky'late blending stocks. The process of this invention is especially designed for the utilization of byproducts to produce also pure toluene and low molecular weight olens and diolens, specically isobutylene and butadiene.

One object of this invention is to provide an integrated process adapted for the catalytic conversion of hydrocarbon oils to aviation base stock of required low acid heat, high aviation octane number and high 'lead response. A further object of this invention is to augment the yield of such aviation base stocks by an improved alkylation process in which selected portions of the low molecular Weight and normally gaseous hydrocarbon products of 3 to 5 carbon atoms per molecule are converted to additional high quality aviation base stocks. A further object of this invention is to provide a process in which the butenes are separated from the low molecular weight hydrocarbon conversion products prior to this alkylation reaction and in which the isobutylene and the n-butenes are thus separately made available as valuable by-Products which are suitable for the preparation of polymers and co-polymers vincluding motor fuels and high molecular weight synthetic rubber-like products, as well as for conversion to'materials also useful in the preparation of such products; for example, the n-butenes may be dehydrogenated to produce butadiene. It is a still further object of this invention to provide a process in which toluene of nitration grade (suitable for the preparation of explosives) is separated from catalytically converted aviation base stock. v

Other objects and advantages of the invention will be apparent from the following description, the drawings and the claims.`

Referring to the drawings, Figure I presents a diagrammatic flow plan of the process of this invention and indicates the flow of materials between the various reaction and treating units involved. Figures II,` IIa and IIb present a more detailed illustration o f suitable types of reaction and treating units which may be used in an apparatus for carrying out the process of this invention. Figure III is a detailed ow sheet indicating quantities and distributionV of the various streams in an example of the invention.

Referring to Figure II, the reference character I designates a charge line through which the oil to be processed is introduced into the system.

40 I oxide-alumina, beryllia-'silica beryllia-alumina.

While this oil is preferably a clean condensate stock, such as a gas oil, it is also possible to employ virgin gasolines, kerosene or crudes or topped or reduced crudes containing unvaporizable material as the charging material. The charge passes through a preheating and vaporizing furnace II wherein the oil is heated to a temperature sufcient to vaporize a substantial portion thereof. The products from the heating furnace may be passed into a separator I2 in which unvaporized materials are separated from the vapors, these being removed overhead by line I3 and passed into a cracking chamber I4 into which an active cracking catalyst of the type later to be described may be introduced through line I5.

In many cases, lthe preheating furnace II and separator I2 may be omitted and the oil may be passed in relatively cold condition through lines I0, I6 and I3 directly into the reactor I4. In this case, the heat necessary for carrying out the cracking operation may be supplied entirely by the hot catalyst introduced into the-cracking chamber through line I5. f y 'While this catalyst may comprise any suitable material for the 'catalytic cracking of hydrocarbon oils, suchy as activated clays and the like, n

the catalyst used is preferably a synthetic gel having as its principal constituents silica and` alumina. Suc'h synthetic gel catalysts may contain from about 5% to 30% alumina, preferably between 10 and 20%', and may be prepared by adding alumina or a solution of an aluminum salt which may be subsequently converted to aluminaI to a lhydrous silica or to a solution from which the hydrous silica is formed. The resulting catalyst may be activated by maintain.. ing it at a temperature from 800 to 1000 F. for a period o1 several hours. Other highly active synthetic gel catalysts may also be used, and may comprise silica-zircona, boria-alumina, tungsten ponent catalysts may be employed containing the form of a vertical upright vessel having an inverted conical base in which the oil to be treated isdischarged.l A perforated grid is preferably provided in the. lower portion of the reaction' chamber I4 forming 'a' distributing plate to insure more uniform'distribution of theoil vapors throughout the catalyst mass. lThe velocity of the oil vapors passing upwardly through the cracking chamber I4 should be adjusted to 's maintain a relatively dense, turbulent iluidized mass of loil vapors and catalysts therein, and may `range, for example, between 0.5 to 5 feet per second te maintain a density of about 5 to 40 pounds per cubic'foot in the lower portion of the cracking chamber.' The velocity ofthe vapors is preferably not so high as to cause entrainment of any substantialproportionof catalyst in the vapors reaching the topv of the cracking chamber. The amount of catalyst introduced into the cracking chamber may. range from 0.5 to 25 or more parts' by ,weight perrpart of oil supplied thereto,

depending upon whether it is desired to supply' all the heat of the cracking operation by the hot catalyst introduced. The time of residence of the oil vapors in the cracking zone in contact .with the catalyst should lbe suillcient to obtain a lconversion of between 40% and 80% of the feed oil to other constituents and may be, for

' all of which may be withdrawn from the system F. `is rich in aromatics and isoparaiilns and conor recycled to the cracking chamber I4 by line 24.

h In accordance with the preferred embodiment `of 'the invention, the naphtha in line 23 constitutes the initial feed fer the production of aviation base fuel by a further catalytictreatment later to be described.- This naphtha-may be a 400 F. or higher Aend point motor fuel fraction, a 300 F.,4 or a 325 F. or 335 F.- end point aviation gasoline fraction, or it may be a light naphl' tha boiling up. to 200 or 225 F. `It has been v found, for example, that the fraction of the' cracked product boiling between ,200 an'd 300A tains only a relatively small amount of oleilns. Hence it maybe withdrawn and used directly asia motor fuel or'aviation`g'a'soline blending y stock. or subjected to extraction .to recover aro-v matics such as toluene contained therein.

A naphthar produced by a single stagecracking treatment of heavier oils is relatively low in octane number and lead response as compared to the naphtha obtained by subjecting this product to Aa second stage catalytic reforming treatment. This second stage ltreatment not only increases the yield and availability of pure toluene in the' product but also so increases the octane vnumber and lead response of this product- (with or without toluene extraction) Vthat the catalytically reformed naphtha is a, valuable blending stock for 'the preparation of 100 octane number aviation gasoline. It is thus preferable in the process of this invention to remove in line 23 a naphtha boiling up to'about 300 to 400 F. and tp subject example, between about 5 seconds and 1 minute A or more,-depending upon the activity of the catalvst. While-the pressure is ordinarily about atmospheric in the cracking chamberfmild superatmospheric pressures of about 2 to 10 latmospheres may be employed.

It will be understood that in accordance with 'the present invention the crackingconditions in the cracking chamber I4 are preferably controlled to' obtain a maximum yield-0f olefinic, isoparailinic and 'aromatic hydrocarbons.' The temperature in the cracking chamber I4 may be maintained between about 750 F. and 11003 F., and preferably between about 875 F, and 1000 F., with the synthetic gel catalyst describedV above.

The cracked products are withdrawn through line I 1 and passed to suitable fractionating equipment. such as towers I8, I9 and 20, operated to produce an uncondensed gas fraction containing this naphtha, to further catalytic treatment.

As will be discussed later the second treating step saturates the olefins. Consequently, Cs hydrocarbons may be'included in the feed in line 23 in accordance with' the requirements for isopentane for the aviation. gasoline vapor pressure.

- In general it is preferableto exclude the Cs cut to the second stage because it is more advantageously converted to alkylate insubsequent al- Vkylation than to aviation base stock in the second stage treating step.

'lhisnaphtha is passed by line 25 through preheater I`524 or by-pass 26 into a second catalytic treating chamber 21, which is suitably of the same design as the cracking chamber I4, and is supplied with a similar powdered catalyst by line 28. Chamber 21 is also operated similarly to chamber I4 in that the velocity of the oil vapors passing upwardly through the treating chamber is preferably controlled to maintain a dense,

agitated, fluid mass of catalyst and oil vapors Y' within the main body of the treating chamber and to avoid carrying overhead with the vapors mainly hydrogen and hydrocarbons of the methane and ethane series which is withdrawn by line 2I, a fraction containing hydrocarbons of 3 to 5V carbon atoms per molecule withdrawn by line 22, and a naphtha fraction withdrawn by line 23. Towers I9Qand 20 are preferably operated at pressures substantially above that of cracking chamber I4 in order to avoid refrigeration'in the renux eonaensers.. This is secured lay/pumpl Isn and compressor I5I, operating respectively on the liquid and gaseous portions of the tower I8Y dis- .Y tillate. The heavier 'materials may also be separated in tower I8 into a residual fraction and gas oil and heavy naphtha fractions, either or any substantial amounts of the catalyst. This catalyst isl preferably of the same general type as that usedin the cracking chamber I4, as this catalyst has been found to be particularly suitable for producing aviation gasoline ofA low acid heat, high aviation octane number and lead response. Because of the large quantity of alkvlate raw materials normally produced in the cracking step, additional aviation base stock for blending therewith may be produced by adding thermally cracked naphthas, thermally reformed naphthas and other olenic gasolines or 'virgin gasolines, kerosenesand light gas oils to thevfeed to reactor 21 through line |53.

The treatingvconditions, such as temperature, catalyst ratio, time of contact, etc. maintained in chamber 21 are adjusted to produce an aviation gasoline blending stock of high aviation octane number and lead response and of low unsaturates and acid heat, one of the principal functions of this second treatment being to reduce the unsaturates content and the acid heat and to'increase the octane number of the aviation fraction of the naphtha feed. In the case of high sulfur distiilates, this further treatment materially reduces the sulfur content ofthe nished product. The temperature maintained in the treating chamber 21 should be between about 600 and 1000 F., and is preferably between about 750 and 900 F. vThe time of contact and the catalyst-oil rratio in chamber 21 may be substantially the same or less than in chamber I4.

While both reactors i4 and 21 are shown to 4be fed from a common catalyst source. separate catalyst streams from separate regenerators of the same or different catalysts may .be used.

'Ihe treated naphtha products are withdrawn by line 29 and passed to suitable fractionating equipment, such as towers 30, 3i and 32, which are operated to produce an uncondensed gas fraction containing mainly hydrogen and hydrocarbons of the methane and ethane series which is withdrawn in line 33, and a fraction containinghydrocarbons of 3 to 5 calbon atoms per molecule which is withdrawnqby line 34. The distillate from tower 30 is also preferably compressed before fractionation, bythe same method used `to compress the distillate from tower IB. It is also desirable to subject this distillate, or the naphtha fraction thereof, to clay treating or other suitable treatment to remove any diolens, peroxides or other unstable materials which may be present. This improves the octance number of the gasoline, and also the operation of the phenol extraction process, as it removes materials which would degrade the phenol and cause troublesome deposits. The feed to tower 32, for example, may be rvaporized in heater |54 and passed at approximately the temperature of its dew point, downwardly over activated clay in Vessel |55, then through fractionating tower 1,56 from which high boiling polymers are removed as bottoms and the distillate is passed into tower 32 for further fractionation. A light naphtha fraction .boiling up to about 175 to 200 F. is .withdrawn by line 35 and is suitable for use as an aviation gasoline blending stock. An intermediate naphtha fraction boiling up to about 300 F, is withdrawn by line 36. 'I'his fraction contains a high proportionof toluene and very little unsaturates. It provides an especially desirable source of pure toluene which may be extracted by a process to be later described. A heavier fraction boiling between about 300 and 4001.1?. may be withdrawn by line 31 and used as a' motor fuel blending stock. This may also be recycled to either catalytic treating unit in order to increase the total yield of toluene. A gas oil and a residual fraction consisting mainly of polymers and entrained catalysts are withdrawn by lines 38 and |51 and either may be recycled to the treating chamber 21 by line 25 or to chamber I4 by line 24.

Considering further the treating chambers I4 and 21, the catalysts contained therein gradually accumulate carbonaceous deposits which reduce their activity. As a result, it is necessary to regenerate these catalysts. 'Ihis may be done moest conveniently .by continuously withdrawing catalyst from the reaction chambers, regenerating it by treatment 'with an oxidizing gas to remove the carbonaceous material, and then returning the regenerated catalyst to the reaction chambers. Where different catalysts are4 used in the two chambers I4 and 21, such regenerative treatment must of course be -conducted in separate units. In the preferred case illustrated. the same catalyst is used in both chambers and is passed through a common regenerator 39,

which may suitably be of the same construction.

as the reaction vessels I4 and 21 as regards the middle and lower portions The catalyst is withdrawn from chamber i4 through a conduit 40, which preferably has anl extended portion 4i projecting upwardly into the reactor. A stripping and iluidizing gas may be introduced into the conduit l40 at one or more spaced points by means of lines "42. This is the preferred methodused in transferring the powdered catalyst in order to maintain it in a mobile or readily flowing state, and it is to be understood that suitable fiuidizing gases (preferably inert or not reactive with the medium' in contact with the catalyst at the particular point of supply) are supplied continuously to all vessels and conduits in which the catalysts may collect. The catalyst from conduit 40 discharges into a stream` of air entering through line 43 and is suspended therein and passed upwardly through line 44 into the regenerating chamber 39. A large amount of heat is .liberated in this regeneration and it is desirable to control the temperature of the When an aviation gasoline of 90% at 293 F. 1

is to be produced as a final product, it is advantageous to include hydrocarbons boiling in the 3D0-335 boiling range in the treated aviationbase, as indicated in Figure III. Then, the low boiling alkylate will give suflicient volatility to the final blend. By doing so the overall octane number and yield will be increased by the inclusion of lthis high octane number (10U-110 with 4 cc. TEL) cut.

bers I4 and 21.

regenerator 39 below a predetermined point which would tend to impair permanently the activity of the catalyst. This control may be accomplished by the amount of catalyst circulated through the reaction chambers or by the l gree of regeneration in chamber 39 is withdrawn through conduit 45 which connects with lines vI5 and 28 and is then returned to the reaction cham- The spentI regenerating gas is withdrawn through line 46 and may be subjected to any suitable treatment for recovery and return of catalyst powder entrained therein.

Returning to' thecracked products, the fraction containing hydrocarbons of 3 to 5 carbon vatoms in lines 22 and 34 contains large proportions of oleflns and isoparalns, thus making it extremely valuable to combine this catalytic cracking process with processes for the recovery and use of these materials. It has been found especially advantageous to separate the isobutylene and normal butenes from this fraction and to subject the remainder to an aikylation treatment to producev increased yields of aviation gasoline.v Of course, where either the isobutylene or the normal butenes are not desired for other purposes,

' veniently as the Ca-Ct cuts) are combined in line 41 and passed through fractionating towers 48 of about 40 to 55%.

illustrated method is to absorb the isobutylene by mixing with cold, concentrated sulfuric acid supplied'by line il, the mixture of sulfuric acid. and C4 cut being'passed through one or more f absorptionvessels' 55 in which it is maintained Y for a time suillcient to permitsubstantially coml,

plete absorption of theisobutylene. The sulfuric acid is preferably supplied at a ooncentra-I 290 F. and a top temperature of about 190 F. The distillate therefrom will also contain aqueous alcohol and polymer, which vare separated from the n-butenes, the alcohol being returned to the feed line to the stripper 6|.

The n-butenes, after washing with aqueous lalkali and water to remove traces of acid, are

then ready for conversion. to butadiene. This may be accomplished yby thermal or catalytic dehydrogenation, preferably conducted at-a low partial pressure of butene of. about 40 to 100mm.

mercury absolute pressure which may be provided by operation under suitable vacuum or by the use of diluent, inert gases or steam', depending upon hydrogenation may be conducted with a finely powdered catalyst suspended in the reacting gases, using. for example, apparatus similar to thatdescribed above for the catalytic cracking tion of about 65% and the absorption temperature is maintained throughout at about 60 F. tov

65 F. in order to absorb the isobutylene selectively without substantial absorption of the normal butenes. the pressure being suillciently high a tower 51 to liberate the isobutylene. This tower is suitably supplied with open steam (which i `tract layer is heated and stripped with steam in may come from subsequent reconcentration of the acid) and may be maintained at a temperature of about 240 F. at the bottom and about 170 F. at thetop. is further fractionated to separate isobutylene from aqueous alcohol'and small amounts of polymer. The aqueous 'alcohol may. be mixed with the acid extract and thus recycled continuously to the stripper wherein it is converted largely to isobutylene. The acid leaving the bottom of the stripper byline 59 will have anacid concentration It may be passed through an acid concentrator 60 wherein 'it is con'centrated to the desired strength of about 65% for l recycling to the absorber.

The distillate from this tower' the type of catalyst employed. The catalytic deoperation, or the catalyst may be held stationari'ly j in a reaction vessel and the reagents passed therethrough. The catalyst usually becomes contami- -nated with carbonaceous materials and may be to maintain the C4 cut in liquid' phase. After f'xperiodically regenerated by treatment with oxi-v the acid and C4 cut are passedthrough one or 1 .more absorption vessels 55 (either concurrently dizing gases, such as air or steam. When using stationary catalyst beds. it is preferred' to main-14 tain two or more. reaction vessels in parallel to 'permit alternate on-stream and regeneration `treatments in each vessel without interrupting the vcontinuous supply of butadiene. This process is illustrated in the drawings. In all types of reaction. it isdesiredA to subject the n-butenes to the dehydrogenation temperature fora very short time of about one-,tenth to one second. Where the nature of the catalyst permits, it is I desirable to heat the n-butenes only to a tem- Since it is necessary to separate isobutylenc from the butenesbefore their recovery. the isobutylene extraction is, necessary whether or not the isobutylenes are to be used as a purified product as in production ,of Butyl rubber. Where it is not used separately, the isobutylene is fed in linev 53 with the Ca and Cs oleilns. Also, the'bottom acid layer from 56 may be heated to 225-275 F. and cooled so as to form the dimer and trimer of lsobutylene which may be fed to alkylation through line i3. For alkylation, this separation from acid is preferred.

The upper layer of unextracted' C4 cutl in separator 56 is withdrawn by line 58 and may be passed directly to the alkylation treatment, or, if desired, it may be treated with a, more concentrated sulfuric acid suitably of 80 to 85% strength, supplied by line |58,-to absorbthe normal butenes. This absorption system may be substantially similar to that just described and may contain one or more absorption vessels I 59, a separator i60, a stripping -tower 6| and a concentrator 62 for recycled acid. 'I'he absorptionA is` suitably also conducted with the C4 cut in liquid .phase at about the same temperature of 60 F. although somewhat. higher temperiatturesy up to'about 100 F. may be used. The stripping tower 6I lis supplied at the bottom with open steam. and is operated to maintain abottom temperature of about perature at which thermal decomposition is not pronounced, say 1100 F., and to mix the heated butenes with highly superheated steam or other gas, which is separately preheated to a temperature sufilciently high, say about 1400 to 2000" F., to bring the mixture up to the desired reactionv temperature of about 1150 to 1300 F. -This involves the use, for example, of about 7 to 12 mols. of steam per mol of hydrocarbon under the temperature conditions just described. It is also desirable to cool the products leaving the dehy- 1 drogenation zone to below a temperature causingV degradation of the products, say below about 1000 F. as quickly as' possible, for examplepby introducing a spray of 4water or cooled reaction products into the hot l,products as they leavethe reaction zone.

Referring to the drawings,the-n0rmal butenes are passed by line :63 through a furnace 64 wherein they .are heated to a temperature just below that at which thermal decomposition becomes l appreciable, for example, about 1100 to 1200 F.,

and are passed .throughline 65 to dehydrogenation vessels-.661m 61. Highly superheated steam is supplied'by line 68 and the resulting mixture at reaction temperature is passed through a bed of a suitable catalyst 69.4 A spray of water is in-` troduced. just below the catalyst bed by line 10 to quench the reaction products which vare removed by line 'Il and passed to suitable fractionation equipment to 'separate a. butadiene product and .unconverted butenes, which may be recycled, from lighter and heavier products. l

The catalyst 69 may be any suitable dehydrogenation catalyst that permits the use oi steam in the reaction vessel and is preferably a catalyst having both dehydrogenating and water gas activity, in order to promote the steam-carbon reaction. Such catalysts, for example, mayv con- 9 tain a major proportion of magnesium oxide, sa about 50 to 90%, lesser proportions of manganese dioxide, chromium oxide or iron oxide .up to about 40% to 49%, and small amounts, between about 0.5% and 20% and generally below 5 or 10% of alkali or alkaline earth oxides,'such as the oxides of sodium andy potassium. Similar small amounts of stabilizers may also be included to prolong the life of the catalyst, such as ever, undergo loss in activity on contact with steam, and hence any substantial amount of water vapor should be avoided when they are used. They may be regenerated by treatment with air, while steam is preferably used for `regeneration of the previously described catalysts.

Such regeneration may be carried out, for example, in dehydrogenation vessel B1 by supplying steam or air alone, heated to the reaction' temperature, through line 12. When air is used, the vent gases are separately removed through line 13.

The dehydrogenated product in line 1| may be cooled to about atmospheric temperature with water. supplied through line |12, and passed into a water and tar separator |13. The gases may then be compressed by compressor 14 and passed through cooler 15 into separator 16 wherein uncondensed gases containing mainly hydrogen and hydrocarbons of less than- 44 carbon atoms per molecule are separated from the liquid hydrocarbons. These uncondensed gases may be further treated by oil'scrubbing or other suitable methods to recover any C4 hydrocarbons which may be passed with the liquid from separator 16 to fractionating towers 11 and 18, hydrocarbons of` less than 4 carbon atoms being withdrawn through line 19 and hydrocarbons of 5 or more carbon atoms being' withdrawn through line 80. The hydrocarbons of 4 carbon atoms per molecule are passed by line 8| to suitable absorption equipment for separating butadiene from butenes, illustrated diagrammatically by a butadiene absorber 82 and stripper 83 through which a suitable solvent, such as a copper salt solution, is circulated by pump 84. Butenes may be recycled to the dehydrogenation treatment by line 85. A portion of the recycled butenes may be passed to the butene extraction treatmentby line |51, in order to prevent build-up of isobutene or butanes in the recycle stream. Concentrated butadiene is withdrawn through line 86, It may be scrubbed with water to remove solvent materials, such as ammonia, and subjected to any further purication that may be desired; generally water scrubbing alone is sufflcient to provide a concentrated butadiene product of sufficiently high purity (98% or more) for use in the preparation of synthetic rubber, such as the Buna" type rubbers prepared by emulsion polymerization with styrene or acrylonitrile, or the Butyl type of rubber prepared by low temperature copolymerization with isobutylene.

While any suitable solvent having selective properties and capable of separating butadiene from mono-olehs may be used in the absorber 82, it is generally preferred to use an aqueous solution of a copper. salt,-such as copper chloride or copper acetate and a nitrogen base, such as ammonia or pyridine.

As indicated above, the C3 and C5 fractions are subjected to alkylation, preferably in admixture with the isobutane fraction of the C4 cut. This alkylation treatment is designed to cause reaction of the isoparaiilns with olefns to form satv urated branched chain hydrocarbons boiling in the motor fuel range. This 4occurs when one molecule of an isoparafn, such as isobutane and the isopentanes, combines with one molecule of an olefin, such as propene, the normal butenes, isobutene, the pentenes and isopentenes. The alkylation products are of high anti-knock value and lead response and provide a valuable aviation gasoline blending stock. They may also be used in motor fuels. The alkylation reaction conditions are adjusted so as to substantially avoid reactions between olefins to form unsaturated polymers. This is accomplished by suitable adjustment of the reaction conditions and catalyst activity and especially by the use of a very large excess of isoparaflins to oleflns in the reaction zone, ratios of total isoparain to' total olefin in the reaction zone ranging from about 100/1 to about 400/1 being illustrative. Suitable catalysts for promoting the alkylation reaction include sulfuric acid, fluorsulfonic acid, chlorsulfonic acid, boron fluoride-water, boron fluoride-phosphoric acid, aluminum chloride-hydrogen chloride, sodium chloride-aluminum chlo" ride and similar complexes of aluminum chloride with other salts, also activated clays and the like. The reaction may be conducted in vapor or liquid phase, the latter being generally used, in which event the pressure on the reaction zone should be sufficiently high to maintain the hydrocarbons undergoing treatment substantially completely in the liquid phase, i. e., the pressure on the reaction vessel should be at least equal to the vapor pressure at the reaction tem'- parafns) are supplied by line 53, see Figure IIA,

and are mixed with a large excess of isobutane from line 81, for example, in suiiicient amount to provide a total isoparaiiin/total olen ratio in the mixture of about 200/1. This mixture is passed through cooler 88 and cooled to a temperature between about 35 and 65 F., preferably sufciently low to maintain a temperature in the reaction zone of about 50 F. Strong sulfuric acid, having a concentration of preferably about 97 to 98% and at least `above about 92%, supplied by line 89, is mixed with the cooled feed stock and the mixture is passed by pump into the bottom of the alkylation reaction vessel 9|. This vessel is provided with suitable means for maintaining an intimate mixture (preferably an emulsion) of the acid and hydrocarbon phases, such as the perforated plate |92, which passes the mixture inthe form of numerous jets under high velocity into the main portion of the reaction zone, which is maintained suitably under a pressure` of about pounds per square inch gauge in order to maintainthe hydrocarbons ture of about 50 F. ,5,

The' feed rate should beadjusted to provide a time of contact of the oleiins in the reaction zone sufilciently long to secure reaction of at least the therein in liquid phase at, the reaction temperamajor portion of theolens.- This may range,

. for example, from about 0.5 hour to 1.5 hours or l longer.

The mixture of acid leaving vessel 9| by line 92 is passed to a separator 93, the lower acid layer separating therein being withdrawn by line 94 and recycled by line 95. A portion oi.' this acid is withdrawn fromthe system and replaced by fresh .concentrated acid supplied 'oy line 96 to maintain the desired acid strength in the alkylation vessel. A portion of the mixture in line 92 may also be recycled directly by line 91 if it is desired'to prolong the time of reaction.

'I'he upper, hydrocarbon layer from separator 'sa is washed with man in vessel sa to' remove acid and is passed to column 99 from which a C: cut is separated from heavier hydrocarbons. These are passed to column |00, which is operated to produce a distillate consisting substantially of isobutane, which is recycled by line 81 to the alkylation unit. The butanes in the original C4 cut (after separation of olens) are also conveniently introduced into the alkylation treatment through vessel by line |Il| in order to avoid passing the inert norma-1 butane of this` cut through the alkylation vessel. The normal butane and less volatile hydrocarbons are withdrawn from the bottom of tower |00 and passed to tower |02, where excess C4 and Cs hydrocarbons above volatility requirements of the nal alkylate may be removed as distillate through line |03, the residue from this column passing to tower |04, where an aviation alkylate fraction boiling up to about 300 F., for example, isvsepa.- rated as distillate from higher boiling residual products. The excess C5 hydrocarbons removed through line |03 are essentially isopentane and this 'isopentane stream, preferably after separation of n-butane, may be recirculated in line 81 with the isobutane so that sufficient isopentane concentration can be built up in the alkylation reactor to .alkylate this excess isopentane. While this step and the alkylation of the Ca C4 and C5 oleilns have been indicated to be carried out in a single alkylation reactor as indicated above, the amount of isopentanes passing to the alkylation treatment may be controlled by the operation. of columns 20 and'32; some isopentane may be withdrawn in the bottoms from these columns to meet volatility requirements in the gasoline blending operation', The individual olen and isoparafiin alkylationsp-may also be carried out in separate reactor syster'ns..A

As indicated above, the catalytically treated naphtha. contains a substantial proportion of aromaticA hydrocarbons and very little olenic hydrocarbons. The U. S. Army specifications for nitration grade; toluene require a purity of above 99% and also vsubstantial freedom from paraiiln andiol'eiin hydrocarbons. This can be readily extracted inhigh yields from the above-described catalytically treated naphthas.

It is desirable to, select for this extraction rather sharply cut naphtha fractions of narrow boiling range as this has been found to aid the separation of toluene from :the other hydrocar- --bons present. The extraction may be conducted with any suitable selective solvent, for example, liquid sulfur dioxide maybe used between about -20 and 60 F, Apreferred method for separating toluene from such naphtha fractions involves subjecting the vaporized naphtha to treatment under distillation conditions withfa solvent which is characterizedby its ability to reduce the vapor pressure of aromatic hydrocarbons to a substantial extent, while at the same time reducing'the vapor pressure of othervtypes of hydrocarbons to a much smaller degree. Phenol is an illustrative solvent oi this type. The separation of toluene with such a solvent is illustrated 5 in the drawing.

,absorption tower I I3, which is maintained at av temperature above the dew point (at the prevailing pressure) of the mixed hydrocarbon feed. Phenol is supplied to a. higher level of the column at a 'temperature substantially equal to that prevailing in the column at the Vpoint of supply by line IM. Heat is supplied to the bottom of the column by any suitablel means, such as heat exchanger H5, in order to increase the concentrating eect of aromatic hydrocarbons in the extract, and cooling means, such as heat exchanger I|6, are supplied at the top of the column in order to provide reux and to prevent loss of-phenol in the vapor raffinate removed by line II'I. This railnate is also suitable blending` stock for aviation gasoline. f

A suitable column design permitting the required separation provides 30 theoretical plates between the phenol feed line H4 and the bottom of the tower ||3, I6 theoretical plates above the phenol feed to the raffinate withdrawal and a 1:1 by volume (cold liquid basis) reflux ratio based on the ralnate. v

The absorption tower |I-3 is preferably operated at a low pressure in order to avoid excessiveheat requirements and may be operated, for example, at a pressure of about 5 to 20 pounds gauge with a .bottom temperature of about 385 F. and a top temperature of about 230 F. Optimum temperatures will be found to vary to some extent in relation to the composition and boiling range of the feed stock and the operating pressures. The phenolic extract is passed by line' f H8 to a stripping or distillation tower ||9 from.'

which the extracted toluene is remove as distillate by line |20 and the phenol as residue by line |2I. Thisphenol may be recycled by line carbon feed (including any recycle extract fractions) supplied thereto. Other ratios of phenol to hydrocarbon, for example in the approximate' range of 0.8 to 6 vols/vol. of hydrocarbon feed. maybeused. 1 ,Y f

The extract distillate in line |20V contains' a high proportion of toluene but requires further purication and refractionationv to -eliminate i small. amounts of solvent and of olens; parafiins and other aromatic hydrocarbons that are generally present. It may,A for. example, be treated vwith sulfuric acid at about 80".F.` (about 25 pounds 0f 98% sulfuric acid per barrel being reference to Figure III of the drawings.

lusuallysuillcient) and then lbe washed in the acid treating and washing steps indicated diagrammatically as |22 and |23 respectively. The treated extract is then redistilled in tower |24 to produce a pure toluene fraction of nitration grade and boiling at 231 F. with'a total deviation o1'. plus or minus 1 F., which is withdrawn by line |25. In the usual operation, where the distillation oi the catalytically treated naphtha fraction prior to the absorption treatment is sufilciently sharp, toluene of this purity will be obtained as the overhead-distillate from tower |24, which will then be operated under the usual conditions with pump-back of a part of the distillate for reflux. If the initial fractionation of the treated naphtha is not sufiiciently sharp, small amounts of low boiling products may be present in the extract 'feed to tower |24; in this event, such low boiling fractions may be separatcd as overhead distillate and the pure toluene product is withdrawn as a sidestream by line |26 from the upper portion of the column several plates below the top. In order to secure this toluene distillate or sidestream fraction of the requisite purity, it is desirable to remove a substantial portion, for example, about 50% of the towerfeed as bottoms therefrom. These contain a high concentration of toluene and are recycled through linev |21 and vaporizer |28 to the absorption tower I I3. These bottoms in line |21 may also be redistilled to remove small amounts of higher boiling fractions or a portion of the bottoms may be removed by line |28 in order to prevent the accumulation of such hi-gh boiling fractions in the system.

A suitable method for conducting the processes of this invention will now be described, with This process, as described, is intended-to be conducted in the equipment illustrated in Figures II, IIa and IIb of the drawings, but it may be conducted in any other suitable type of equipment for carryingA out the various treating steps and reactions which are involved.

Referring to Figure III, the quantities of the various streams are indicated on the ow plan as barrels per day. Hydrocarbons are indicated by the number of carbon atoms per molecule and olens by the symbol indicated by iso C4=, propene by Cs=, pentenes by and normal butenes by n-C4=. The charge to the catalytic cracking unit is a distillate gas oil stock of 31.5 A, P. I. gravity, 179 F. aniline point and 650 F. midboiling point. The cracking operation is conducted with a powdered synthetic silica-alumina gel cracking catalyst with an average temperature in the cracking zone of 975 F. and a ratio of 5.8 pounds of catalyst per pound of oil in the feed to the cracking chamber, the oil being passed through this chamber at a rate of 3.0 poundsper hour Thus, isobutene is per pound of catalyst present in the cracking chamber.

The naphtha cut (containing hydrocarbons from 6 carbon atoms per molecule up to those boiling at 400 F.) has a 49.7 A. P. I. gravity and a midboiling point of 250 F. It is passed to the catalytic naphtha treating unit in which it is subjected to an` averaged reaction temperature of 900 F., using a powdered synthetic silica-alumina gel catalyst with a ratio of 8 pounds of catalyst per pound of oil in the feed to the treating chamber, the oil being passed through this chamberat a rate of 1.3 pounds per hour per'pound of catalyst present therein. The light ends from 14 both thecatalytic cracking and treating units, including hydrocarbons of 4 Acarbon atoms per molecule, are passed through light end recovery equipment in which the Cis and most of the Css are separated from the lighter gases. A portion of the C; cut, containing propene is passed to the alkylation unit. The C4 cut is passed under suificient pressure to maintain it in liquid condition, through an isobutene extraction unit, which is maintained at a temperature of F. and is.

supplied with sulfuric acid of strength in a proportion of about 1 to 2 mols of acid per mol oi' isobutylene in the hydrocarbons treated. The hydrocarbons remain in the extractor for about 50 minutes. The acid extract is' diluted to an acid concentration of 45% and/is stripped with steam, thus obtaining isobutene and some polymers, equivalent to 89 barrels/day -of Cis.

The rafnate is passed through a normal butene extractor maintained at a temperature of F. and under suilcient pressure to maintain the hydrocarbons in liquid phase, at a rate permitting about three hours time of residence of the hydrocarbon in the extraction vessel. Sul- Hiuric acid of 83% strength is supplied to this extractor in a proportion of about 1 mol per mol of butene in the feed hydrocarbons, the acid extract being diluted to an acid concentration of 62% in the regenerator or stripper in which noriron oxide, 5% potassium oxide and 5% copper oxide, unconverted normal butenes being separated from the product and recycled.

The raiinate from the butene extraction, polymer from the isobutene extraction, and the Ca cut from the catalytic cracking and treating units are passed to an alkylationunit. Isobutene maiT also be passed to this unit or may be used separately for the prepa-ration of synthetic rubber. Isobutene polymers and copolymers formed in the isobutene and butene extraction steps may also be passed to the alkylationunit or may be depolymerized to yield additional isobutene,'if desired. C5 cut from the catalytic cracking units is also supplied to the alkylation unit in an amount suilcient to make up the olefin requirements of this process. This is about 1 mol of total olens (including those available by depolymerization of the isobutene polymers) per mol of isobutane supplied thereto. The alkylation reaction vessel is operated at a temperature of 50 F. and at a pressure suicient to maintain the hydrocarbons therein in liquid condition. It is supplied with sulfuric acid in a ratio of about 1 volume per volume of hydrocarbon feed. The hydrocarbon feed is passed through the reaction vessel at a rate to provide a time of residence therein of about 30minutes and in suitable proportions to maintain a ratioof volumes of isobutane per volume of olen therein. Unreacted hydrocarbons, particularly isobutane, are recycled to provide a ratio of about 5.5 volumes of isobutane per volume of olefin in the feed mixture supplied to the alkylation reactor.

' 15 The intermediate naphtha fraction of the product from the catalytic naphtha treating unit, boiling between about 175 and 255 F. is

passed to the toluene extraction plant, preferably after treating with clay to remove any suitable cleiinic or peroxide materials which may be i present and would react with the phenol. The

clay treatment may be conducted by passing the naphtha cut in vapor phase over lAttapulgus clay of 30 to 60 mesh, for example, at apressure of about 100 pounds per square inch and at a tem-A perature of about 400 F., or to be more exact, at aboutthe dew pointof the material at the treating conditions. The naphtha feed rate is about 2.3 barrels (measured as cold liquid) per hour per ton of clay in the treating zone. Higher boiling polymers formed in this treatment are separated from the naphtha out, which is then passed to the toluene extraction plant.

The toluene is suitably extracted by being scrubbed with phenol in an extraction vessel in which the naphtha'is supplied atabout 15 pounds per square inch gauge pressure and at a temperature just about its dew point under-these conditions, the phenol being supplied-at a temperature of about 250 F. in a proportion of about 3 volumes per volume of oil feed (both measured as cold liquid). I

The following products may thus be obtained from a gas oil feed of 14,4000 barrels per day:

Butadiene standard tous per year 22,600 sobutene B /D 1-- 414V Butane 'R /D 292 C4 polymers....v. .B./D...A 87 Pentane (surplus above aviation gasoline 1 requirements) B./D Tolene (nitration grade) .BJD... Aviation gasoline 2 (7 lb. Reid vapor pressure, 86.3 ASTM clearv octane number) B ,/D-- For motor gasoline blending lB./D Gas oil R /D 1 E uivalent'to 18,800 long tons er year of butyl rubber. 2 Tgis, blended with 818 B./D. o? 69 O. N. natural naphthe, gives 5496 B./D. of 100 octane aviation gasoline with 4 cc. TEL per gallon added.

While the invention has been described with particular reference to various types of apparatus and treating agents in each of the process steps described herein, the invention should not be so limited, for these have been presented as illustrative embodiments of the many variations that are readily apparent as being included within the scope of this invention.

We claim: 1. In a process for the catalytic conversion of n relatively heavier hydrocarbon oilsto relatively more volatile motor fuels and gaseous products,

' an improved method of utilizing said products comprising separating therefrom a fraction comprising mainly olen and parain hydrocarbons of 3 to 5l carbon atoms per molecule, separating a mixture `of hydrocarbons of 4 carbon atoms per inolecule from said fraction, then removing. isobutene from this mixture vby selective extraction -with sulfuric acid of a strength linsuiiicient to react with the normal .butene present, heating isobutene, removing normal butenes from the re'- mainder of said niixture, then subjecting a mixture of the. remainder of said first mentioned fraction, the remainder` of said fraction of 4 carcontent permissible 'in'aviation fuel, separating said naphtha traction, subjecting said naphtha fraction to\ `further treatment in the presence of a synthetic gel catalyst containingV silica and alumina, to reduce `the amount of unsaturates contained therein, separating an aviation gasoline fraction from the products ofsaid second catalytic treatment, extracting toluene from said aviation gasoline fraction, also-separating from the products of said cracking treatments a fraction comprisingl mainly olefin and paraffin *hydrocarbonsl or 3 to 5 carbon atoms per molecule, separating ai; least a part oi' the olefins of 4 carbon atoms per molecule from said CFC fractions, and sub- `iecting the remainder of thisfraction to alkylation conditions adapted tol cause reaction of the olefins with-'the isoparafilns present therein. lto

produce an 'alkylate product comprising saturated isoparafiins, and -mixing said alkylate product with the''esidue of saidaviation gasoline product aftersaid toluene extraction to yield an improved" aviation grade gasoline of high octane number and low in unsaturates and in acid heat. 3. Process according to claim 2 in whichthe olefins of 4 carbon atoms are removed from the said hydrocarbon fraction of 3 to 5 carbon atoms per molecule before it is subjected to the said alkylation treatment.

4. Process accordingto claim 2 in which the normal butenes are removed from the said hydrocarbon fraction oi 3 to 5 carbon atoms per molecule before it is subjected to the said alkylation treatment.

5. Process for producing aviation gasoline, comprising cracking a higher boiling oil in the presence of a cracking catalyst to produce relatively lower boiling motor fuel and normally gaseous products containing substantial proportionsof olefins, separating from the cracked products a naphtha fraction boiling in the range comprising hydrocarbons from 6 carbon atoms per .molecule at least up to those boiling at 200 F. and below 400 F. and containing olens in excess of the amount permissible in aviation gasoline, then subjecting said` naphtha fraction to a further treatment in the presence of a reforming catalyst under catalytic reforming conditions to reduce the amount of oleflns therein, separating an aviation gasoline fraction from the products of said second catalytic treatment, alsoseparating from the products of at least the first of said catalytic i the acidextract and separating therefrom a-mixture of isobutene polymers comprising mainly ditreatments n v andfparaiiin hydrocarbons of 3 to 5 carbon atoms per molecule, separating olensof 4 carbon atoms.' 'per molecule from saidfraction under conditions wherein vthe said separated oleilns of 4 carbon atoms per molecule comprisesubstantially all of said olens present in' said Ca-s fraction, subjecting the remainder of said fraction to alkylation conditions adapted to cause olens and iso-paraffins present therein to form saturated iso-parafbon atoms per molecule comprising n -butane and isobutane and the said-isobutene polymers to fins of 7 to 9 carbon atoms per molecule and of good antiknock characteristics and blending the product or alkyianonwith the aviation gasoune a fraction consisting mainly of olen 17 fraction, polymerizing the separated isobutene to a low molecular weight polymer comprising mainly di-isobutene, said polymer being passed to said up to those boiling at 200 F., said naphtha fraction having an olen content in excess of that permissible in aviation fuel, also separating from the cracking product a fraction comprising mainly oleiin and parailln hydrocarbons ot 3 to 5 carbon atoms per molecule, separating substantially all olens of 4 carbon atoms per molecule from said fraction, then subjecting the remainder of said fraction to alkylation conditions adapted to cause reaction of the oleilns of 3 and 5 carbon atoms with the iso-parafiins present therein, to form saturated iso-paraillns of 7 to 9 carbon atoms per molecule and having good antiknock characteristics, treating the said naphtha fraction to remove excess oleflns and nt the naphtha fraction for .use as an aviation fuel, as an incident of such treating, producing aw fraction comprising mainly olefin and paradln hydrocarbons of 3 to 5 carbon atoms per'molecule. subjecting said less mentioned fraction to alkylation with admixture with the first mentioned fraction comprising mainly olefin and paraffin hydrocarbons of 3 to 5 carbon atoms per molecule, blending the resulting alkylate with said naphtha fraction which has been treated to remove excess oleiins, extracting toluene from the naphtha fraction,

blending the remainder thereof with the alkylated product to yield an improved fuel of high antiknock value and relatively low acid heat.

ROBERT P. RUSSELL. EGER V. MURPHREE.

' CHARLES E. HEMMINGER.

REFERENCES CITED UNITED STATES PATENTS Number I Name Date 2,178,602 Morrell Nov. 7, 1939 l5 2,203,829 Lener Jan. 11, 1940 2,211,747 Goidsby Aug. 13, 1940 2,240,134 Eglof! Apr. 29, 1941 2,256,615 Hederhorst Sept. 23, 1941 2,259,723 Ballard Oct. 21, 1941 2,264,447 Meier Dec.'2, 1941 2,276,171 Ewell Mar. 10, 1942 2,279,547 Zimmerman Apr. 14, 1942 2,283,851, Day May i9, 1942 2,297,773 Kanhofer Oct. 6, 1942 2,304,183 Layng Dec. 8, 1942 2,310,327 Sweeney Feb. 9, 1943 2,328,754 Thomas Sept. 7, 1943 2,328,773 Benedict Sept. 7, 1943 2,335,596 Marschner Nov. 30, 1943 2,371,355 4 Ross et al. Mar. 13, 1945 OTHER REFERENCES Preprint, High-Octane Aviation Fuel by the Sulphuric Acid Alkylation Process," Twentieth Annual Meeting, A. P. I., November 17, '1939,

pages 1 to 12.

Oil and Gas Jour. article, Technique for retining war products explained; March 19, 1942, pages 18 and 19. 

